Process for obtaining 2-hydroxy-4-methylthiobutyric acid (MHA)

ABSTRACT

A method for the isolation of 2-hydroxy-4-methylthiobutyric acid (MHA), wherein MHA is isolated from a reaction mixture obtained by addition of hydrogen cyanide (HCN) to methylmercaptopropionaldehyde (MMP) and hydrolysis by sulphuric acid of the methylmercaptopropionaldehyde cyanohydrin (MMP-CH) thus obtained. The reaction mixture is brought into contact in a liquid/liquid extraction system with an organic solvent substantially immiscible with water, in order to form an extraction solution which contains the solvent and the MHA transferred out of the reaction mixture. By bringing the salt content of the reaction mixture, prior to the liquid/liquid extraction, to a concentration of about &gt;50 wt. % (wt./wt.), preferably &gt;55 wt. %, referred to the sum of the inorganic constituents of the reaction mixture, the coordinated use of energy in the total system is improved, the evaporation of strongly corrosive solutions is avoided, the efficiency of the hydrolysis step is increased and the distribution coefficients during the extraction are improved. The MHA is isolated as the extract from this extraction solution by evaporation, and may be used as animal feed supplement.

FIELD OF THE INVENTION

The invention relates to a method for the isolation of2-hydroxy-4-methylthiobutyric acid (MHA), whereby MHA is isolated from areaction mixture obtained by addition of hydrogen cyanide (HCN) tomethylmercaptopropionaldehyde (MMP) and hydrolysis by sulphuric acid ofthe methylmercaptopropionaldehyde cyanohydrin (MMP-CH) thus obtained,the reaction mixture being brought into contact in a liquid/liquidextraction system with an organic solvent substantially immiscible withwater, in order to form an extraction solution which contains thesolvent and the MHA transferred out of the reaction mixture, and the MHAis isolated as the extract from this extraction solution by evaporation.

2-hydroxy-4-methylthiobutyric acid (MHA) is the hydroxy analogue of theessential amino acid methionine in racemic form and, like this acid, isan important additive in animal nutrition. In the rearing of poultry MHAexhibits growth-stimulating properties similar to those of the aminoacids known for this. This additive is also becoming of increasinginterest in other areas of animal nutrition.

MHA is mostly used in the form of aqueous concentrates containing, inaddition to the monomer, a certain proportion of oligomers, mainly thedimeric and trimeric linear ester acids. The content of these oligomersis dependent on the conditions of preparation and on the concentrationchosen. Owing to their lower nutritive efficiency and the unfavourableinfluence on the flow properties as a result of an increase inviscosity, it is however desirable to maintain their percentage as lowas possible. Commercially available formulations having a totalconcentration of 88 to 90 wt. % contain up to 24 wt. %, corresponding toabout 27 mol %, in total of oligomers, corresponding to a ratio ofmonomer to oligomers of about 3:1.

The general method for the preparation of MHA proceeds from3-methylpropionaldehyde, also referred to asmethylmercaptopropionaldehyde or MMP, which is reacted with hydrogencyanide to form 2-hydroxy-4-methylthiobutyronitrile, also referred to asMMP cyanohydrin or MMP-CH (equation I). ##STR1##

The MMP cyanohydrin formed is then hydrolysed, generally by strongmineral acids such as sulphuric acid or hydrochloric acid, via theintermediate step involving the formation of2-hydroxy-4-methylthiobutyramide, also referred to as MHA amide(equation II) ##STR2## to prepare the methionine hydroxy analogue (MHA)(equation III). ##STR3##

This hydrolysis may be carried out either in one step or in two steps.Here, by "steps" it is meant that mineral acid and/or water is addedeither once or twice in order to hydrolyse the MMP-CH, that is, thenumber of steps corresponds to the number of addition procedures.

BACKGROUND OF THE INVENTION

The following publications are cited as being close prior art:

EP-A-0 142 488=D1,

EP-A-0 143 100=D2,

EP-A-0 330 527=D3 and

WO-A-94/28717 = EP 93 924 374=D4.

A similar type of method for the isolation of MHA is known, for example,from D1. D1 employs two-step hydrolysis using sulphuric acid in order toisolate MHA in liquid form as a highly-concentrated aqueous solution.

According to D1, MHA is obtained after the hydrolysis reaction, which iscarried out via the amide step using excess mineral acid under specifiedconditions of concentration and temperature, by means of a solventextraction wherein use is made of certain solvents which are partiallymiscible with water.

According to the information in D1, the characterising feature of themethod described in that document is to be seen in the isolation of MHAfrom the extraction solution, which is carried out in such a way thatthe isolation includes the removal of the organic solvent in thepresence of at least about 5 wt. % of water, referred to the remainingextract (MHA). MHA is isolated from the extraction solution bydistillation (see Examples), with steam distillation being preferred. Asa result of removing the solvent from the extraction solution during thesteam distillation, the discharge obtained is a mixture of MHA andwater. The steam distillation is therefore carried out in such a waythat the discharge contains at least 5 wt. % of water.

At another place in the text of D1 it is stated that the columnconditions during the distillation are controlled in such a way thateverywhere in the column, but at least in the bottom fraction, theliquid phase contains 5 wt. % of water.

It follows from this that in the absence of a sufficient quantity ofwater during the isolation of MHA from the extraction solution, theincreasing formation possibly of undesirable by-products (dimers andoligomers) is to be expected.

Furthermore, in the distillation the steam serves as an operative agentfor the complete removal of the extracting agent from the MHA solution,for example, through the formation of a low-boiling azeotropic mixturewith the respective extracting agent.

Further embodiments of a method essentially similar in type aredescribed in D2. In contrast to the preamble in D1, the hydrolysis ofthe MMP-CH by a mineral acid is described, with reference being made tothe alternative possibility of using HCl instead of H₂ SO₄. Altogetherthree additional variants are disclosed, which are concerned essentiallywith variations in the method of isolating MHA from the mineral acidhydrolysate or in the working up of the extract during the liquid/liquidextraction.

In an initial aspect according to D2, the hydrolysate is brought intocontact with the organic solvent without previous separation of anyessential fractions from solid substances contained therein. Moreover,according to D2 the conditions of the extraction are controlled so thatthe extract and the aqueous raffinate are the only liquid phases, whichare formed during phase separation after the extraction.

A disadvantage of this first aspect is that the extraction is laden withthe whole of the salt component formed in the hydrolysate duringhydrolysis, which leads to a relatively high mass flow of hydrolysateand accordingly also of solvent. This results in correspondingly highenergy costs in the solvent evaporation and condensation and costs ofcorresponding loss of solvent and of correspondingly large units forextraction and evaporation. A lowering of the operating and investmentcosts at this point in the process would therefore be desirable(especially in view of the size of such a plant and the potential foreconomy associated therewith).

The raffinate obtained from the extraction as a homogeneous liquid phasemust, according to D1 or D2 or D4, be freed from remains of solvent bystripping or distillation, which is an undesirable additional expense.

In a second aspect, D2 refers to the separation of the organic solventfrom the extract. On this point, it is stated that the separation iseffected by subjecting the extract to a steam distillation, with thesolvent being distilled off and a bottom fraction of aqueous MHA beingformed.

A disadvantage of the use of steam is in particular the increasedaccumulation of aqueous solvent-laden steam condensate, which hassubsequently to be freed from solvent at undesirable additional expense,such as that of distillation or stripping, in order then to return it tothe process at a suitable point, or otherwise it has to be expensivelydisposed of, for example, by burning. Avoidance of additional strippingsteam would therefore be desirable.

Finally, in a third aspect, D2 lays emphasis on the nature of thesolvent to be used for the liquid/liquid extraction. The criteria to beconsidered in the selection of a suitable solvent include in particularthe following points:

the boiling point of the solvent is to be between 60° C. and 200° C.;

the distribution coefficient for MHA in equilibrium between hydrolysateand solvent is to be at least approximately two;

the distribution coefficient of the solvent in equilibrium betweenextract and aqueous phase is to be at least approximately one;

the solubility of water in the solvent at room temperature is to be notmore than about 12 wt. %.

The relatively high boiling point range, of from 60 to 200° C., of thesolvents to be used here necessitates for the evaporation of the extractrelatively elevated temperatures, which may impair the product, as wellas additional auxiliaries, such as stripping steam, which isundesirable.

One of the considerable disadvantages of the methods in D1 and D2consists however in the high salt content, which forms during thesaponification and which nevertheless unfavourably contaminates themethod for the isolation of MHA in an otherwise relatively elegantliquid/liquid extraction process. A working up of the inevitablyresulting mixture of ammonium salts is in most cases not economic, thedisposal is very hazardous from the environmental aspect and in theforeseeable future is likely to be prohibited by law even at siteshaving few rigid conditions.

There has been no lack of attempts to lessen or even to avoid theaccumulation of salt from the saponification, but the advantages gainedthereby were in every case achieved by accepting a number of otherdisadvantages or by dispensing with the elegant handling according to D1and D2.

Thus in D3 there is disclosed a single-step method of hydrolysis usingsulphuric acid as saponifying agent, which is carried out withoutsolvent and leads directly to concentrated aqueous MHA solutions, withcrystalline ammonium sulphate in marketable form being obtained as thecoproduct. This object is achieved by neutralising the saponificationmixture with ammonium hydroxide solution to the extent that the excessmineral acid and the ammonium bisulphate formed are converted into theneutral sulphate with the formation of two liquid phases, which fortheir part are separated and evaporated in order to isolate firstlyliquid MHA and secondly crystalline ammonium sulphate. During this thedifferent filtration and recirculation steps are combined in such a waythat virtually no product is lost and no waste water contaminated withsalt is formed. The resulting MHA is of a quality similar to that of theproduct obtained in D1.

However, even this relatively environmentally harmless method hasvarious disadvantages. When reproducing this method, the Applicant ofthe present invention found that firstly, because of the comparativelyhigh dilution of the sulphuric acid (20-50%), excesses of aciddefinitely higher than those given have to be used in order to achieve acomplete cyanohydrin conversion. Moreover, to avoid precipitations ofsalt during neutralisation the method has to be carried out at higherdilution, to render possible a clean separation of the two liquidphases. Secondly, the ammonium sulphate isolated has a stickyconsistency and is tainted with a strong smell, so that anaftertreatment such as, for example, a washing filtration orrecrystallisation appears unavoidable and the method is therebyadditionally made more expensive. Moreover the method--unlike what ispostulated--consumes more energy in the evaporation steps than does themethod cited by way of comparison in D1. Further, it is cost-intensiveand very expensive as regards apparatus for the treatment of solidsusing filtration/centrifugation, which involves two separate paths, aswell as the drying of the ammonium sulphate, which is not shown in theflow diagram.

A partial solution to the dilemma is promised by D4. D4 discloses therecovery of sulphuric acid from a sulphate-containing flow of wastematerial which arises during the preparation of2-hydroxy-4-(methylthio)butyric acid by hydrolysis of2-hydroxy-4-(methylthio)butyronitrile using sulphuric acid.

The recovery of sulphuric acid from ammonium sulphate, ammoniumbisulphate and/or residues containing sulphuric acid has for a long timebeen prior art in the preparation of MMA and, just as is known for theresidues from the saponification of acetone cyanohydrin, is achieved bycombustion in a so-called split-contact plant of the flows of wastematter arising during saponification and extraction.

Here, in a manner familiar to the person skilled in the art, SO₂ isfirst of all produced as a decomposition product, which is oxidised onthe contact catalyst to form SO₃, which is finally converted intosulphuric acid. The resulting sulphuric acid can then be returned againto the saponification process, while the other former constituents ofthe "load of salt" may be found substantially in the form of combustiongases.

Elegant as this method may be, it is also not free from disadvantages.Thus in the methods according to D1 and D2 flows of waste materialarise, whereof the sulphate concentration is relatively low, butinvariably too low to permit direct introduction into a split-contactplant. Hence a concentration or increase by mixing with concentratedflows of waste material from other processes is generally essential.Conventional solutions employed for the operation of split-contactplants have a sulphate salt content of >50 wt. %. Higher concentrationsare even more preferred. Concentration by evaporation of the waste waterarising from the isolation of MHA is however, owing to the highcorrosiveness of the waste water, a relatively costly undertaking which,from the selection of special materials for the evaporation equipment tothe special safety precautions required, is excessive and expensive.

SUMMARY OF THE INVENTION

In the light of the prior art cited and discussed here and of thedisadvantages associated with the known methods, it is the object ofthis invention to provide another method for the preparation of2-hydroxy-4-methylthiobutyric acid (MHA) of the type mentioned at thebeginning, which is to be as simple and economical as possible asregards the working up of the reaction products and is to permit ashighly concentrated a product as possible having as low as possible acontent of dimers, oligomers and by-products. The new method is, ifpossible, to retain the advantages of the simple practicability of thestep involving isolation of MHA in a liquid/liquid extraction, but is atthe same time to permit as simple and rational as possible a disposal ofthe load of salt accumulating in the sulphate-containing waste water. Inparticular, a process is to be provided which inter alia permits thedirect introduction of immediate flows of waste water into, for example,a split-contact plant, for the recovery of sulphuric acid which isusable and hence can be recirculated in the process.

This object and others not stated in detail are fulfilled by a method ofthe kind described at the beginning, which possesses the features of thecharacterising part of claim 1.

Advantageous variants of the method are placed under protection in themethod claims dependent on claim 1.

By bringing the salt content of the reaction mixture, prior to theliquid/liquid extraction, to a concentration of about >50 wt. %,preferably >55 wt. % (wt./wt.), in each case referred to the sum of theinorganic constituents of the reaction mixture, according to theinvention a method is provided which permits the preparation of MHA ofoutstanding quality and which at the same time in a not readilyforeseeable manner solves, or at least surprisingly improves, theproblem of the inevitably accumulating salt. In particular, there arefurthermore a number of considerable advantages in increasing the saltconcentration in a suitable manner after the hydrolysis but before theliquid/liquid extraction and hence never immediately prior to thepossible connection to a split-contact plant.

The advantages mentioned include the following.

The evaporation otherwise appropriate prior to the introduction into asplit-contact plant can be completely omitted.

More simply constructed and hence cheaper equipment is adequate for aconcentration prior to the extraction because, owing to its particularcomposition, the concentrated solution is very much less aggressive atthis particular point in the procedure and is therefore, in particular,also less corrosive.

In the case of at least partial shifting of the evaporation to aparticular earlier point in the procedure there is an overall improvedcoordinated use of energy. Since already tempered (hot) hydrolysate isevaporated, and not raffinate which has already been cooled byextraction, the energy consumption is less.

In the course of a concentration by evaporation, the separation ofunwanted low-boiling components of the hydrolysate is improved.

The hydrolysis of the MHA amide can be carried out in dilute solution,which results in a more complete chemical conversion. As it is better,in the hydrolysis of MHA amide using sulphuric acid, to employ a moredilute sulphuric acid (<40 wt. %) in order to allow the hydrolysis totake place as completely as possible, a less highly concentratedsulphuric acid can be used in the hydrolysis of the MHA amide, withoutits being necessary to anticipate that the subsequent extraction with anorganic solvent will proceed less favourably. Owing to theconcentration, a deterioration of the distribution coefficient isavoided and, in particular, less MHA remains in the raffinate during theextraction.

The method according to the invention therefore also fulfils inparticular the requirement for an advantageous sulphuric acidconcentration during the saponification and for the isolation of MHAfrom the hydrolysate in association with the provision of a raffinatewhich, owing to its composition, is more suitable for working up in asplit-contact plant, while at the same time the energy balance overallis considerably improved.

The concentration of the salt in the hydrolysate of about >50 wt. % is aquite useable concentration for the subsequent working up by means of asplit-contact plant. Ranges of from 55 to 60 wt. % are preferred.Particularly preferably the salt content of the reaction mixture isadjusted to about 60 to 80 wt. %, in each case referred to the sum ofthe inorganic constituents of the reaction mixture. To determine thisvalue, which may also be termed the content referred to the"organic-free basis", essentially the water content, the sulphuric acidcontent and the content of sulphate ions and ammonium ions are used.These are the main inorganic constituents of the hydrolysate.

The term "concentration" includes, for the purpose of the inventiongenerally, the increase in the salt concentration (referred to a basisfree from organic constituents in wt. %="organic-free basis").

In a preferred variant of the method according to the invention, for thepurpose of concentration a suitable quantity of ammonium sulphate isadded to the reaction mixture (hydrolysate) which has formed as a resultof the addition of HCN to MMP and hydrolysis by H₂ SO₄ of the MMP-CHformed. The concentration of salt is accordingly increased by anaddition. This variant in particular has a number of great advantages.

As already mentioned, in principle two mutually opposing effects have tobe considered. On the one hand, the hydrolysis reaction should proceedas far as possible to completion, to which end a relatively lower amideconcentration in the water present during the MHA amide saponificationis appropriate. This inevitably results in a dilute solution of MHA andammonium hydrogen sulphate. On the other hand, for the extractionprocess it is more advantageous to decrease the water content of thehydrolysate, that is, to have as high an MHA concentration as possiblein the aqueous phase.

In D1 and D2 the hydrolysis is therefore carried out at an approximately<40 wt. % sulphuric acid concentration, which inevitably leads howeverto a dilute solution of MHA and ammonium hydrogen sulphate. To improvethe subsequent extraction, in D1 and D2 the ammonium hydrogen sulphateis converted into neutral ammonium sulphate by the addition of anhydrousammonia. This improves the corrosion behaviour of the solution but itcan lead to the precipitation of solid substances, which may adverselyaffect the operation of an extraction. Accordingly, in D1 and D2 watermay subsequently be readded, in order to bring the precipitated saltsback into solution. In the course of this the concentration of MHA orsalt should not however be lowered excessively, as otherwise extractionmay be impeded.

In contrast to this, it is possible by the addition according to theinvention of ammonium sulphate to carry out the hydrolysis of the MHAamide in more dilute solution, which results in a more completeconversion during the hydrolysis step. At the same time, the solutionneed not be neutralised. That is, the treatment with anhydrous ammoniaand a possible redilution are unnecessary.

In a preferred embodiment, the method of the invention is carried out insuch a way that, prior to isolation of MHA, ammonium sulphate is addedin a quantity effective for salting out, with two phases being formed.

It is known from D1 and D2 that the presence of a high saltconcentration (most advantageously ammonium hydrogen sulphate) "saltsout" MHA and consequently has a favourable effect on the distributioncoefficients. However, according to D1 and D2 a two-phase system is tobe avoided, as this would interfere with the extraction. In contrastthereto, it has been found that a two-phase system formed by addingammonium sulphate is particularly beneficial to the extraction step andaltogether has a positive effect on this. Thus, for example, in onevariant, solvent is added in such a quantity to a two-phase mixtureobtained according to the invention, that two distinct phases are formedand a separation need be carried out only once.

Moreover, the addition of ammonium sulphate at this point of the methodfor isolating MHA has the general advantage of sparing the product.Through the addition of ammonium sulphate only the ammonium saltconcentration is increased and not simultaneously the MHA concentration.An additional thermal stress (discoloration) of the target product aimedfor need not occur; on the contrary, owing to the dissolving of theammonium sulphate there is a lowering of the temperature. Neverthelessthe distribution coefficient is favourably affected.

Although the variant of the ammonium sulphate addition discussed abovepossesses indisputable advantages in the sparing of the product, in analternative variation of the method according to the invention it may bepreferred to increase the salt concentration of the reaction mixture(hydrolysate) by evaporation.

Depending on the salt concentration in the raffinate, there may resultincreased discharge of solvent at the bottom of the column which,according to the publications D1, D2 or D4, has to be separated off in araffinate stripping stage. In contrast to this, in the present inventionit was ascertained that a solvent recovered in this way, due to athermal stress in strongly acid medium is highly contaminated withby-products and is consequently unsuitable for direct recirculation intothe extraction system.

Surprisingly, it has become apparent within the scope of the inventionthat even slight cooling of the product in the bottom of the columnresults in the discharge of a two-phase mixture of aqueous and organicraffinate, with the organic raffinate consisting to the extent of >97%of solvent, which immediately after a simple separation in a separatingvessel can be passed back into the extraction system without furtheradditional expense and the loss of solvent can in this way be minimised.

Consequently, a particularly preferred variant of the method accordingto the invention, wherein the salt concentration is increased byevaporation, is characterised in that at least three liquid phasesimmediately result from the extraction system.

Here in principle it is made possible for a homogeneous raffinate and anextract consisting of two liquid phases to be formed, the first liquidphase in the extract consisting substantially of MHA, solvent and smallportions of water, while the second liquid phase consists substantiallyof water, MHA and small portions of salt, but it is far moreadvantageous that, in a variant of the method according to theinvention, a homogeneous extract and a raffinate consisting of twoliquid phases be formed. In this case, it is particularly advantageousthat the first liquid phase in the raffinate consist substantially ofammonium salt and water and of small portions of MHA and organicsolvent, while the second liquid phase consists substantially of organicsolvent and of small portions of water and MHA.

In an advantageous variation of the method of the invention, the stepsfollowing the actual concentration by evaporation are conducted in sucha way that the second liquid phase contains MHA in a quantity of from0.01 to 0.5 wt. %, contains solvent in a quantity of from 90 to 99 wt. %and contains water in a quantity of from 0.1 to 10 wt. %, while thefirst liquid phase contains water in a quantity of from 20 to 50 wt. %,contains MHA in a quantity of from 0.01 to 0.5 wt. % and contains saltin a quantity of from 50 to 80 wt. %; the constituents of each phasetaken separately must total 100 wt. %.

The extraction solution employed for the concentration by evaporation inorder to isolate MHA is recovered from the reaction mixture byextraction. In principle, one may of course use all organic solventsknown in prior art which exhibit a number of the properties alreadyindicated above in the introduction to this description. The organicsolvent employed for the extraction should be substantially immisciblewith water. A partial miscibility of the organic solvent with water ishowever tolerable. The solvents which are suitable for the separation ofsubstances in the liquid/liquid extraction include a large number whichmeet the conditions of chemical inertness and a low capacity to dissolvewater. In general it is preferred that the solubility of water in thesolvent at room temperature be not more than about 15 wt. %, preferablynot more than 10 wt. %. Of the suitable solvents, those having a boilingpoint of between about 60° C. and about 200° C., preferably of betweenabout 70° C. and 150° C., are preferred. The distribution coefficientbetween the solvent containing the extracted MHA, and the aqueousraffinate remaining behind after the contacting of the solvent and theMHA hydrolysate should be at least about 2 for MHA in equilibrium. Thisdistribution coefficient is preferably at least 5. Moreover thedistribution coefficient for MHA in equilibrium between extractionsolution and washing water should be not less than about 1.0. Inaddition, the solvent is to exhibit a low toxicity.

A number of ketones, aldehydes and carboxylic esters are particularlysuitable as solvents for the extraction. Particularly preferred solventsare ketones of relatively low molecular weight such as, for example,methyl n-propyl ketone, methyl ethyl ketone, methyl amyl ketone, methylisoamyl ketone, methyl isobutyl ketone, ethyl butyl ketone anddiisobutyl ketone. Aldehydes such as, for example, n-butyraldehyde, andesters such as, for example, ethyl acetate, n-butyl acetate, n-propylacetate and isopropyl acetate are also suitable solvents for theextraction. Alcohols may also be used, although these are less preferredowing to their mutual solubility with water, a slow phase separation andthe tendency to react with MHA.

Compared with these solvents already used or proposed in prior art,which are of course also of a certain usefulness for the presentinvention, totally unexpectedly it has been found that the use of ethercompounds as solvents in the extraction is associated with manyadvantages. The ethers which may be used according to the inventioninclude primarily those corresponding to the general formula I

    R.sup.1 --O--R.sup.2                                       (I),

wherein R¹ and R² independently of one another are identical ordifferent C₁ -C₅ alkyl, linear or branched. Suitable ether compoundsinclude the following:

    ______________________________________                                        No.       R1          R2       Bp [° C.]                               ______________________________________                                        1         Ethyl       Ethyl    35                                               2 n-Propyl Methyl                                                             3 n-Propyl Ethyl                                                              4 n-Propyl n-Propyl 90                                                        5 i-Propyl i-Propyl 69                                                        6 n-Butyl Methyl 71                                                           7 n-Butyl Ethyl 92                                                            8 n-Butyl n-Propyl                                                            9 n-Butyl n-Butyl 143                                                         10 tert. Butyl Methyl 56                                                      11 tert. Butyl Ethyl 73                                                       12 tert. Butyl n-Propyl                                                       13 tert. Butyl n-Butyl                                                        14 Neopentyl Methyl                                                           15 Neopentyl Ethyl                                                            16 Neopentyl n-Propyl                                                         17 Neopentyl n-Butyl                                                        ______________________________________                                    

The preferred ether compounds are firstly those which show no or only alow tendency to form peroxides such as, for example, MTBE. Asymmetricalethers are also preferred.

Secondly, those compounds which have a boiling point of <60° C. are veryuseful, as they can be removed from the target product completely andeasily.

Most especially preferred within the scope of the invention is the useof methyl tertiary butyl ether (MTBE), which meets all theabove-mentioned criteria.

The actual extraction can in principle be carried out continuously orintermittently. An agitated tank, for example, is suitable for abatch-operated procedure. Preferably, however, the extraction is carriedout in a continuous countercurrent extraction plant possessing anextraction zone designed for accelerating the mass transfer betweensolvent and aqueous phase. Thus it is advantageous, for example, tocarry out the extraction in a cascade of continuous countercurrentmixer-settlers, a packed column, a perforated-plate column, preferablyin the form of a pulsed column or column having moving plates, arotary-disk extractor or a centrifugal extractor. In a particularlypreferred embodiment, the extraction is carried out in aperforated-plate column for liquid/liquid extraction. Intermittent orpulsed flows, although cyclical and therefore not continuous for thepurpose of rapid flow rates, are regarded as being "continuous" inconnection with the present disclosure.

The extraction process is preferably controlled so as to set up andmaintain the solvent phase as the continuous phase in the extractionzone.

The extract is if necessary washed with water in order to decrease to aminimum the salt content of the end product. Within certainconcentration ranges of the resulting extract solutions the washing mayhowever be dispensed with, in particular at sulphate contents of <0.5wt. % and in view of the salt concentration of the raffinate flowingoff, which should as far as possible not be further diluted. In acontinuous countercurrent extraction system, the extract can be washedby being mixed with water at a point upstream, relative to the directionof the flow of organic substances, of the point at which the hydrolysateis introduced into the liquid/liquid extraction system. Thus, forexample, in a vertical column and in the case of a solvent having aspecific weight preferably of less than 1, solvents are introduced intothe column at a point below the inlet point at which aqueous hydrolysatesolution is introduced, and washing water is introduced into the columnat a point above the inlet point of the hydrolysate solution.

The productivity of the extraction process is raised by operating at asomewhat elevated temperature, in order to provide for a relatively lowviscosity of the solvent phase within the extraction system. Operationat a temperature within the range below the boiling point of the organicsolvent used moreover has a barely favourable effect on the MHAdistribution coefficients between the organic and the aqueous phase.

Within the scope of the invention, MHA can be isolated from theextraction solution by evaporation, as mentioned above. That is, thepresent invention is concerned with a further aspect, in particular withthe evaporation of an extraction solution, such as is obtainable from aliquid/liquid extraction of a reaction mixture obtained, for example, byhydrolysis of MMP-CH using sulphuric acid. In this connection theevaporation is preferably carried out in such a way that the remainingextract contains less than 4 wt. %, preferably less than 2 wt. %, ofwater. Here, in a not readily foreseeable way, a highly concentratedliquid MHA containing particularly low proportions of oligomers anddimers is produced. In view of the known prior art, it is more thansurprising that this can be achieved using a smaller proportion of waterthan may be inferred, for example, from the known publications (D1 andD2).

In a particularly preferred embodiment according to the invention, theorganic solvent is removed during evaporation in a unit providing abrief residence time for the extraction solution in one evaporationstep. The organic solvent is particularly preferably separated from theextraction solution during the evaporation by means, therefore, of afalling-film evaporator, film evaporator and/or short-path evaporator orwith the aid of such a unit.

The expression "with the aid of such a unit" means, within the scope ofthe invention, that the aforesaid units providing a brief residence timefor the extraction solution may also be combined with equipment, knownto the person skilled in the art, for the separation of the solvent fromextraction solutions. Here the units used for the combination need notnecessarily be of the kind having a brief residence time. At this pointone may mention, inter alia, distillation columns which mayalternatively also be equipped for the introduction of steam or othersuitable stripping agents. Combinations including several of the unitsmentioned having a brief residence time are also feasible.

In an advantageous variation of the method according to the invention,it is preferred that the evaporation of the extraction solution becarried out in such a way that as low a content of residual solvent aspossible results. This is achieved, for example, by a combination ofseveral of the above-mentioned units with a stripping stage, which maybe contained as an additional unit or integrated into theabove-mentioned units in the evaporator system, for example, by thedirect introduction of the stripping medium into such an evaporator.

The specific conditions for the evaporation necessarily vary accordingto the particular solvent selected for use in the extraction. Inprinciple, for the evaporation employing a separating unit providing abrief residence time for the extraction solution, it is preferred thatthe pressure during the evaporation be not more than 600 mbar,preferably not more than 400 mbar and particularly preferably not morethan 200 mbar.

The temperature applied during the evaporation is as a rule likewisedependent on the solvent being separated. It is however intended, andtherefore within the scope of the invention is also particularlypreferred, that the temperature during the evaporation be not higherthan 150° C. If this temperature is very definitely exceeded, there maybe resulting thermal damage to the product aimed for. In thisconnection, the temperature during the evaporation is not necessarily tobe understood as the contact temperature of the product with the surfaceof the evaporator unit equipped for brief contact with the product. Thetemperature during the evaporation means rather the average temperaturein the evaporator unit. The temperature at the surface of the evaporatorunit can in case of doubt be very much higher than 150° C. The brevityof the contact time in the evaporator units used is crucial. By thismeans a thermal damage is avoided, even if the contact temperature isdefinitely above 150° C.

Regarding the temperature distribution, within the scope of theinvention it has been found that it is particularly advantageous to theproduct quality if the temperature of the remaining extract immediatelyat the discharge point from the evaporator unit be between 30 and 100°C., preferably from 50 to 95° C. and particularly preferably from 70 to90° C.

As already mentioned, the residence time of the remaining extract isdecisive for the quality and composition of the MHA product aimed for.In an advantageous development of the method according to the invention,the residence time of the remaining extract in the evaporation processis not longer than 1.5 h. This refers to the residence time in theentire evaporator system, which includes at least one evaporation stepwith a very brief residence time. The residence time in the unit havinga very brief residence time, contrary to the maximum of 1.5 h given forthe total residence time, is to be established rather in the range ofminutes or less. In any rate, within the scope of the invention it ispreferred, in the event that the evaporation consists solely of one filmevaporator and/or falling-film evaporator and/or short-path evaporator,that the residence time in these units be not longer than 1 hour andpreferably 40 min.

In a further aspect the method according to the invention, in additionto improving the isolation of MHA from the reaction mixture obtained byhydrolysis using sulphuric acid, also improves the hydrolysis of theactual MMP-CH. Thus in a preferred embodiment according to theinvention, the hydrolysis of MMP-CH is carried out in such a way that,in a first step, MMP-CH is hydrolysed using from 60 to 85 wt. %,preferably from 65 to 80 wt. %, of sulphuric acid in the molar ratio ofMMP-CH to H₂ SO₄ of from 1.0:0.5 to 1:1.0, preferably from 1:0.6 to1:0.95, at temperatures of from 30 to 90° C., preferably from 50 to 70°C., with substantially MHA amide being obtained. Here MHA amide isformed substantially from MMP cyanohydrin, the mixture formed beingmoreover substantially free from unreacted MMP cyanohydrin. In otherwords, this means that the hydrolysis proceeds virtually quantitatively.

A particularly advantageous target product can be isolated by thevariations of the method for recovering MHA described in more detailabove. The improved MHA is characterised according to the invention bycontaining more than 95 wt. % of total MHA as the sum of monomeric MHA,MHA dimers and MHA oligomers (=total MHA) and by having a water contentof between more than 0.1 wt. % and less than 5 wt. %. In particular ithas proved advantageous to the invention that, without greater losses ofquality, an MHA can be obtained which is characterised in that itcontains more than 98 wt. % of MHA as the sum of monomeric MHA, MHAdimers and MHA oligomers and has a water content of between 0.1 wt. %and less than 2 wt. % and a kinematic viscosity of >100 m² /s at 25° C.In this connection it has surprisingly been found that the kinematicviscosity, measured using a Cannon-Fenske viscometer, of highconcentrate (that is, MHA having a content of active components of atleast 98 wt. %), after storage and dilution is comparable with thekinematic viscosity of an 88 wt. % product. Notwithstanding a relativelyhigh content of dimers and oligomers of about 50 wt. %, which isestablished in the high concentrate after storage for about 300 days atroom temperature, when the stored high concentrate has been diluted withwater to about 88 wt. % its kinematic viscosity corresponds to that ofthe 88 wt. % commercial product, which in parallel storage experimentshad an equilibrium concentration of only about 25 wt. % of dimers andoligomers. The equilibrium state was achieved in both cases, both in thediluted high concentrate and in the commercial product. This fact isvery surprising and proves to be a great advantage of a highlyconcentrated MHA variant prepared according to the invention. In view ofthe fact that dimeric and oligomeric constituents of MHA in generalinterfere in a manner disadvantageous to practical processing, it wasall the more surprising that, in spite of high starting contents in theso-called high concentrate, an easily pumpable and thereforetransportable mixture having a favourable viscosity can be obtained.This has various advantages: in particular the viscosity and above allthe high content of active components result in the high concentratebeing transportable more economically, as less water is transported; atthe point of destination in the feed mill the high concentrate can stillbe diluted with water to the usual commercial concentrations, withoutunfavourable higher viscosities having to be accepted.

It has also been found within the scope of the invention that MHA ofparticularly high quality can be obtained by suitably conducting thehydrolysis reaction in combination with the mild evaporation and verybrief residence time to be used according to the invention. Thisparticularly advantageously obtainable MHA is characterised primarily bya content of the sum of dimers and oligomers, referred to the total sumof MHA, of <10 mol %, preferably of <7 mol %. This means that, contraryto the prejudice widely held in prior art, it is possible to obtain ahighly concentrated MHA which, owing to an extremely low proportion ofdimers and oligomers, is a very suitable form of short-term transport.For a longer period of transport it is then preferred, by the additionof water and application of an elevated temperature, to reconvert intomonomeric MHA the dimers and oligomers which increasingly form dependingon the period of storage.

It is also possible within the scope of the invention to use thehighly-concentrated MHA product for the preparation of animal feedsupplements. Here it has been found that all nutritive useful materialsbasically required by the market can be prepared, without loss ofquality, by mixing the MHA concentrate with water, methionine and/orsalts of MHA (ammonium-MHA is preferred) (optionally NH₃ in order toproduce NH₄ -MHA).

First and foremost, in carrying out the invention it has been found thatthe mixtures are obtainable not only by addition of suitable componentsof the mixture such as water, methionine and/or ammonium-MHA from theoutlet of the target MHA product from the evaporation step, but thatlikewise and moreover particularly advantageously in the case of mixingwith ammonium-MHA, ammonia can be introduced directly into the MHAproduct from the evaporation. Here, depending on the quantity of ammoniaadded, a required proportion of MHA is converted into ammonium-MHA.

BRIEF DESCRIPTION OF THE DRAWINGS

The invention is further explained below with reference to the attachedFigures. The following procedures are illustrated by the Figures.

FIG. 1 shows a flow diagram for the isolation of MHA by salt separationand liquid/liquid phase separation in the MHA hydrolysate; theprocedures 1), 2) and 3) can be followed independently of one another;

FIG. 2 likewise shows a flow diagram of an embodiment of the methodaccording to the invention, wherein a salt separation withoutliquid/liquid phase separation is provided; here, too, the procedures 1)and 2) are independent of one another;

FIG. 3 shows a flow diagram of another embodiment according to theinvention, whereby MHA is isolated without salt separation;

FIG. 4 shows a flow diagram of a further embodiment of the invention,whereby MHA is isolated after the salt content has been increased; and

FIG. 5 shows a diagram of a suitable apparatus for carrying out themethod of the invention;

FIG. 6 shows a diagram of another suitable apparatus for carrying outthe method of the invention.

DETAILED DESCRIPTION OF THE INVENTION

In a variant of the method shown in FIG. 1, MHP cyanohydrin (MMP-CH) isconverted into the acid hydroxy analogue of methionine (MHA) in atwo-step hydrolysis reaction using aqueous sulphuric acid. The primaryMHA hydrolysate formed is then evaporated, starting from a concentrationof <40 wt. % of MHA, to a concentration of >40 wt. %, preferably >45 wt.% of MHA, so that two liquid phases are formed.

The water obtained during the evaporation is condensed and returned tothe hydrolysis step with the condensation temperature, in order to saveenergy, being maintained as close as possible to the temperature atwhich the hydrolysis takes place. The fraction of malodorous low-boilingcomponents obtained is to a large extent separated by the steam, removedat the top, optionally with the aid of stripping gases, for example air,and is preferably passed without previous condensation directly to acombustion furnace. The latter may also be a component part of a plantfor the recovery of sulphuric acid (a so-called split-contact plant).

The two liquid phases obtained from the bottom of the evaporation unitare separated from one another at a temperature which exceeds roomtemperature, but is at highest the temperature of the evaporation.

The lower aqueous phase, containing mainly the ammonium salt formed, iscooled until a considerable portion of the dissolved salts crystallisesout. (Procedure 1) or 2)). The temperature required for this is below30° C. The salt crystallisate obtained is separated from the supernatantsolution by centrifugation or filtration. The salt crystallisate may bewashed with a suitable organic solvent, or even with water or an aqueoussalt solution, in order to remove useful material (MHA) still adhering.

The upper organic phase, containing mainly MHA as well as the aqueousfiltrate and possibly the organic filtrate, is separated or, afterpartial or complete prior mixing, is together passed to a liquid/liquidextraction system (procedure 1) or 2)) and separated by means of anorganic solvent into at least two phases, namely, into at least onemainly organic extract solution containing the solvent and MHA and smallproportions of water and salt, and into an aqueous raffinate, whichconsists mainly of salt and water and which is then preferably passed toa plant for the recovery of sulphuric acid (procedure 1)), andoptionally in addition into an organic raffinate, which consists mainlyof solvent and small portions of MHA or water and which can be returnedto the extraction system.

The organic extract solution is passed to a system for the evaporationof the extract, the evaporated solvent and possibly correspondingportions of water being recovered by condensation and returned to theextraction step. The MHA high concentrate obtained as discharge from thebottom of the evaporation unit is adjusted to the required MHAconcentration, preferably between 78 and 98 wt. %, in a conditioninginvolving addition of required quantities of water and/or appropriateadditives such as, for example, methionine or MHA-NH₄ salt.

The salt crystallisate, after an optionally performed salt wash, can bepassed to a purifying or conditioning step (procedure 1)), whereinmarketable ammonium sulphate is produced by the addition of appropriatequantities of NH₃ and subsequent crystallisation and drying, or else itcan be passed in the form of the unrefined product directly to thedrying unit. The salt crystallisate may also, in particular after beingdissolved in water, be passed as a >60% concentrated solution to a plantfor the recovery of sulphuric acid (procedure 2)). Here it isparticularly advantageous to dissolve the salt crystallisate, stillmoist from filtration, in the raffinate from the extraction step and topass the highly concentrated salt solution obtained, having a saltcontent of >75 wt. %, to the plant for the recovery of sulphuric acid,because a salt content of at least 60 wt. % is necessary for this andmoreover each additional increase in concentration contributes to animprovement of the energy balance of such a plant. The concentration ispossible here, especially in the absence of an energy-intensiveevaporation of the salt solution obtainable from the process. All orpart of the sulphuric acid thus recovered may be returned to the MHAhydrolysis step.

It can also be advantageous to pass the aqueous phase, withoutseparation of salt, directly to a plant for the recovery of sulphuricacid, together with the raffinate from the extraction (procedure 3)).Here, too, it is advantageous that the salt content be definitely above60 wt. %. A loss of about 2.5% of the theoretical yield of MHA occurshere, which is still dissolved in the aqueous phase. A considerableadvantage here, however, is the great easing of the extraction orevaporation step, as the inlet flow to the extraction and hence also theuse of solvent can be almost halved as compared with conventionalmethods (cf. D2), which is associated with an extreme saving in energy,especially as regards the evaporation and condensation of the solvent.

In a two-step hydrolysis reaction represented in FIG. 2, MMP cyanohydrin(MMP-CH) is converted into the acid hydroxy analogue of methionine(MHA), in a two-step hydrolysis reaction using aqueous sulphuric acid.The primary MHA hydrolysate formed is then evaporated, starting from aconcentration of <40 wt. % of MHA, to a concentration of >40 wt. %,preferably >45 wt. % of MHA, so that two liquid phases are formed.

The water obtained during the evaporation is condensed and returned tothe hydrolysis step with the condensation temperature, in order to saveenergy, being maintained as close as possible to the temperature atwhich the hydrolysis takes place. The fraction of malodorous low-boilingcomponents obtained is to a large extent separated by the steam, removedat the top, optionally with the aid of stripping gases, for example air,and is preferably passed without previous condensation directly to acombustion furnace. The latter may also be a component part of a plantfor the recovery of sulphuric acid.

The two liquid phases obtained from the bottom of the evaporation unitare cooled together with one another until a suspension of saltcrystallisate and a homogeneous organic/aqueous liquid phase is formed.Here it is advantageous to cool the phases to room temperature.

The salt crystallisate is separated from the supernatant solution bycentrifugation or filtration. The salt crystallisate is washed with asuitable organic solvent, or even with water or an aqueous saltsolution, in order to remove useful material (MHA) still adhering.

The filtrate and possibly the organic filtrate are separated or, afterpartial or complete prior mixing, are together passed to a liquid/liquidextraction system and separated by means of an organic solvent into atleast two phases, namely, into at least one mainly organic extractsolution containing the solvent and MHA and small proportions of waterand salt, and into an aqueous raffinate, which consists mainly of saltand water and which is then preferably passed to a plant for therecovery of sulphuric acid (procedure 1)).

The organic extract solution is passed to a system for the evaporationof the extract, the evaporated solvent and possibly correspondingportions of water being recovered by condensation and returned to theextraction step. The MHA high concentrate obtained as discharge from thebottom of the evaporation unit is adjusted to the required MHAconcentration, preferably between 78 and 98 wt. %, in a conditioninginvolving addition of required quantities of water and/or appropriateadditives such as, for example, methionine or MHA-NH₄ salt.

The salt crystallisate, after an optionally performed salt wash, can bepassed to a purifying or conditioning step (procedure 1)), whereinmarketable ammonium sulphate is produced by the addition of appropriatequantities of NH₃ and subsequent crystallisation and drying, or else itcan be passed in the form of the unrefined product directly to thedrying unit.

The salt crystallisate may also, in particular after being dissolved inwater, be passed as a >60% concentrated solution to a plant for therecovery of sulphuric acid (procedure 2)). Here it is particularlyadvantageous to dissolve the salt crystallisate, still moist fromfiltration, in the raffinate from the extraction step and to pass thehighly concentrated salt solution obtained, having a salt content of >75wt. %, to the plant for the recovery of sulphuric acid, because a saltcontent of at least 60 wt. % is necessary for this and moreover eachadditional increase in concentration contributes to an improvement ofthe energy balance of such a plant. The concentration is possible here,especially in the absence of an energy-intensive evaporation of the saltsolution obtainable from the process. All or part of the sulphuric acidthus recovered may be returned to the MHA hydrolysis step.

In the variant of the method represented in FIG. 3, MMP cyanohydrin(MMP-CH) is converted into the acid hydroxy analogue of methionine(MHA), in a two-step hydrolysis reaction using aqueous sulphuric acid.The primary MHA hydrolysate formed is then evaporated, starting from aconcentration of <40 wt. % of MHA, to a concentration of >40 wt. %,preferably >45 wt. % of MHA, so that two liquid phases are formed.

The water obtained during the evaporation is condensed and returned tothe hydrolysis step with the condensation temperature, in order to saveenergy, being maintained as close as possible to the temperature atwhich the hydrolysis takes place. The fraction of malodorous low-boilingcomponents obtained is to a large extent separated by the steam, removedat the top, optionally with the aid of stripping gases, for example air,and is preferably passed without previous condensation directly to acombustion furnace. The latter may also be a component part of a plantfor the recovery of sulphuric acid (a so-called split-contact plant).

The two liquid phases obtained from the bottom of the evaporation unitare if necessary cooled together, but only to the extent that no saltcrystallisate is formed.

The product of the evaporation is passed to a liquid/liquid extractionsystem and separated by means of an organic solvent into at least twophases, namely, into at least one mainly organic extract solutioncontaining the solvent and MHA and small proportions of water and salt,and into an aqueous raffinate, which consists mainly of salt and waterand which is then preferably passed to a plant for the recovery ofsulphuric acid. The required salt concentration of at least 60 wt. %depends absolutely essentially on the degree of evaporation of theprimary hydrolysate. In this connection it should however be taken intoaccount that this should be only so great that no salt crystallisatesare formed within the extraction system as a result of an excessivelyhigh concentration. The salt concentrations achievable thereby aretherefore less than those in the procedures shown respectively in FIGS.1 and 2. All or part of the sulphuric acid thus recovered may bereturned to the MHA hydrolysis step.

The organic extract solution is passed to a system for evaporating theextract, the evaporated solvent and possibly corresponding portions ofwater being recovered by condensation and returned to the extractionstep. The MHA high concentrate obtained as discharge from the bottom ofthe evaporation unit is adjusted to the required MHA concentration,preferably between 78 and 98 wt. %, in a conditioning involving additionof required quantities of water and/or appropriate additives such as,for example, methionine or MHA-NH₄ salt.

In the variation represented in FIG. 4, MHP cyanohydrin (MMP-CH) isconverted into the acid hydroxy analogue of methionine (MHA) in atwo-step hydrolysis reaction using aqueous sulphuric acid. The primaryMHA hydrolysate formed, which has a concentration of <40 wt. % of MHA,is then subjected to evaporative cooling, wherein the temperature,starting from a reaction temperature of >100° C., is decreased to asuitable lower temperature, for example, 60° C. and at the same time afraction of malodorous low-boiling components together with smallquantities of steam is separated off by distillation, preferably withapplication of a vacuum and optionally with the aid of stripping gases,for example air, and can be passed without previous condensationdirectly to a combustion furnace. The latter may also be a componentpart of a plant for the recovery of sulphuric acid.

By the subsequent addition of (NH₄)₂ SO₄ and/or NH₄ HSO₄ to thehomogeneous MHA hydrolysate solution, the salt concentration presenttherein is increased until two liquid phases are formed but at the sametime a substantial proportion of undissolved solids is not left behind.

The two liquid phases are separated from one another at a temperaturewhich exceeds room temperature (procedure 1)). The upper organic phase,containing mainly MHA, is passed to a liquid/liquid extraction system(procedure 1)) and separated by means of an organic solvent into atleast two phases, namely, into at least one mainly organic extractsolution containing the solvent and MHA and small proportions of waterand salt, and into an aqueous raffinate, which consists mainly of saltand water.

The raffinate, preferably together with the lower, aqueous phaseobtained from the liquid/liquid phase separation and containing mainlythe ammonium salt formed, is passed to a plant for the recovery ofsulphuric acid (procedure 1)).

A loss of about 2.5% of the theoretical yield of MHA occurs here, whichis still dissolved in the aqueous phase. A substantial advantage here isthe great easing of the extraction or evaporation step, as the inletflow to the extraction and hence also the use of solvent can beconsiderably decreased as compared with conventional methods (cf. D2),which is associated with an extreme saving in energy, especially asregards the evaporation and condensation of the solvent.

The organic extract solution is passed to a system for evaporating theextract, the evaporated solvent and optionally corresponding portions ofwater being recovered by condensation and returned to the extractionstep. The MHA high concentrate obtained as discharge from the bottom ofthe evaporation unit is adjusted to the required MHA concentration,preferably between 78 and 98 wt. %, in a conditioning involving additionof required quantities of water and/or appropriate additives such as,for example, methionine or MHA-NH₄ salt.

Alternatively, the two liquid phases can also together be passed to aliquid/liquid extraction system (procedure 2)). The raffinate obtained,a >60% concentrated salt solution, can be passed directly to a plant forthe recovery of sulphuric acid (procedure 2)), because a salt content ofat least 60 wt. % is necessary for this and moreover each additionalincrease in concentration contributes to an improvement in the energybalance of such a plant. The concentration is possible here, especiallyin the absence of an energy-intensive evaporation of the salt solutionobtainable from the process, which is a great advantage. All or part ofthe sulphuric acid thus recovered may be returned to the MHA hydrolysisstep.

The following examples of preparation further explain the subject matterof the invention.

Analytical methods of determination and definitions

The contents of MMP cyanohydrin, MHA amide and MHA monomer respectivelyin the prepared solutions were determined quantitatively by means ofHPLC by comparison with an external standard (pure substance).

    The content of total MHA=MHA amide (optionally)+MHA(=total MHA) monomer+MHA (dimers+oligomers)

was determined by titrimetric determination of the thioether functionusing KBr/KBrO₃ standard solution and was expressed as the sum of thecorresponding MHA monomer equivalents in [wt. %] or [g] or [mol] or [mol%].

The content of MHA dimers+MHA oligomers (DIM+OLI) was established bycalculating the difference of total MHA less MHA monomer (+optionallyMHA amide) and was expressed as the sum of the corresponding MHA monomerequivalents in [wt. %] or [g] or [mol] or [mol %].

The water content was determined by Karl-Fischer titration, the solventcontent was determined by GC or by subtraction, the sulphate or ammoniumcontent was found by ion chromatography using a standard method and thetotal salt content by converting the sulphate or ammonium contents or bysubtraction.

EXAMPLE 1 Continuous Preparation of MHA Hydrolysate Solution

In a two-stage series of stirred-tank reactors 8.7 kg/h of an MHA amidesolution was produced by a continuous charging of 4.2 kg/h (31.3 mol/h)of 97.7% MMP cyanohydrin and 4.5 kg/h (29.7 mol/h) of 65% aqueous H₂ SO₄at a temperature of 50° C. and with an average total residence time of60 min. The MHA amide solution was further converted to form 12.3 kg/hof MHA hydrolysate solution by means of continuous dilution using 3.6kg/h of water in a two-stage series of stirred-tank reactors with areaction tube connected in tandem at a temperature of 90 to 110° C. andwith an average total residence time of 180 min. The reaction solutionaccumulating initially was evaporated to small bulk by continuousintroduction into an evaporator system at a pressure of 100 mbar and wascooled to a temperature of 50° C. at the discharge point. Thepre-evaporated MHA hydrolysate obtained (10.8 kg/h) had the followinganalytical composition: ##EQU1##

EXAMPLE 2 Preparation of MHA-MTBE Extract Solution Experiment 1

2.5 kg of MHA hydrolysate (43.7 wt. % of total MHA, prepared asdescribed in Example 1) together with 1.5 kg of MTBE (technical) wereplaced in a 5 1 mixing vessel equipped with a bottom discharge valve andstirred intensively for 10 min at room temperature. After stirring hadbeen concluded, the two liquid phases formed were separated from oneanother. The procedure was repeated in total 4 times, each time usingfresh solutions.

The organic phases and the aqueous raffinate phases were each combinedand analysed. The compositions of the phases in [wt. %] are shown inTable 1 below.

                  TABLE 1                                                         ______________________________________                                                 Organic phase    Raffinate phase                                       (13.0 kg)  (7.0 kg)                                                                  [wt. %]          [wt. %]                                             ______________________________________                                        Total MHA          41.8             1.9                                         MHA  38.4  1.9                                                                DIM + OLI  3.4  0                                                             H.sub.2 O  4.8 (calc.) 40.0                                                   MTBE (calc.) 53.0  0.04                                                       SO.sub.4.sup.2-  0.2  47.7                                                    NH.sub.4.sup.+  0.02  9.7                                                   ______________________________________                                    

Experiment 2

Experiment 1 was repeated using 2.5 kg of MHA hydrolysate and 1.5 kg ofMTBE, which had been recovered by evaporating MHA-MTBE extract solution(cf. Example 3). The compositions in [wt. %] may be seen in Table 2below.

                  TABLE 2                                                         ______________________________________                                                 Organic phase    Raffinate phase                                       (2.6 kg)  (1.4 kg)                                                                   [wt. %]          [wt. %]                                             ______________________________________                                        Total MHA          42.0             1.9                                         MHA  37.7  1.9                                                                DIM + OLI  4.3  0                                                             H.sub.2 O  4.8 (calc.) 40.0                                                   MTBE (calc.) 53  0.04                                                         SO.sub.4.sup.2-  0.2  47.7                                                    NH.sub.4.sup.+  0.015  9.7                                                  ______________________________________                                    

EXAMPLE 3 Preparation of MHA High Concentrate

In FIG. 5 is shown a diagram of the arrangement of the apparatus usedfor Example 3. This consists substantially of the following equipment:

    ______________________________________                                        001          storage vessel                                                      002   Sambay evaporators each having 0.06 m.sup.2 exchange                   003     surface and a heated double jacket                                    004  receiver for the MHA product                                              005  condensation system for distilled-off solvent                              each consisting of a water-cooled and a brine-cooled                       006   laboratory cooler, a receiver and a water suction pump                    having an adjustable vacuum                                               ______________________________________                                    

Description of the method with reference to FIG. 5

The MHA-MTBE extract solution leaving the extraction is fed continuouslyfrom the storage vessel 001 into the Sambay evaporator 002, which isheated externally. The discharge from 002 is fed via a needle valve intothe likewise heated Sambay evaporator 003; the MHA product dischargedtherefrom is collected in the receiver 004 and there analysed. Thedistillate consisting of solvent is collected in the receivers of thetwo condensation systems 005 and 006 and from there can be returned tothe extraction unit (cf. Example 2, Experiment 2).

Experiment 3

Use of 0.95 l/h (0.85 kg/h) of MHA-MTBE extract solution from Example 2,Experiment 1

Sambay 002:

pressure 250 mbar

temperatures:

heating jacket 125° C.

discharge point 79° C.

Composition of the MHA high concentrate in the bottom discharge from 002

Total MHA: 98.0 wt. %

H₂ O: 0.5 wt. %

Sambay 003:

pressure 50 mbar

temperatures:

heating jacket 140° C.

discharge point 90° C.

vapours 30° C.

Composition of the MHA high concentrate in the bottom discharge from 003

Total MHA: 99.0 wt. % MHA 83.9 mol % DIM+OLI 16.1 mol %

H₂ O: 0.5 wt. %

MTBE: <10 ppm

0.36 kg/h of MHA high concentrate having the above composition wasobtained from the bottom discharge of the Sambay evaporator 003.

Experiment 4

Use of 0.96 l/h (0.86 kg/h) of MHA-MTBE extract solution from Example 2,Experiment 2

Sambay 002:

pressure 250 mbar

temperatures:

heating jacket 125° C.

discharge point 96° C.

Composition of the MHA high concentrate in the bottom discharge from 002

Total MHA: 98.5 wt. %

H₂ O: 0.9 wt. %

Sambay 003:

pressure 50 mbar

temperatures:

heating jacket 120° C.

discharge point 100° C.

vapours 28° C.

Composition of the MHA high concentrate in the bottom discharge from 003

Total MHA: 100.0 wt. % MHA 85.7 mol % DIM+OLI 14.3 mol %

H₂ O: 0.0 wt. %

MTBE: <1 ppm

0.36 kg/h of MHA high concentrate having the above composition wasobtained from the bottom discharge of the Sambay evaporator 003.

EXAMPLE 4 Recovery of Salt from MHA Hydrolysate Prior to the Extractionby Liquid/Liquid and Liquid/Solid Phase Separation (cf. FIG. 1)Experiment 5 MHA Separation by Liquid/Liquid Phase Separation andLiquid/Liquid Extraction Using MTBE

502 g of MHA hydrolysate containing 43.7 wt. % (219.4 g) of total MHA(prepared as described in Example 1) was evaporated at a pressure of 50mbar to a content of 50 wt. % of total MHA. The concentrate (438.7 g)consisted of two liquid phases, which were separated from one another atT=65° C. The composition of the two phases is given in Table 3.

                  TABLE 3                                                         ______________________________________                                                Organic phase                                                                              Aqueous phase                                              (292.4 g) (143.4 g)                                                         ______________________________________                                        Total MHA 73.0 wt. %     3.9 wt. %                                               = 213.4 g = 5.6 g                                                             = 97.3% of theor. = 2.6% of theor.                                           H.sub.2 O 10.6 wt. % = 31.0 g 20.6 wt. % = 29.5 g                             Salt (calc.) 16.4 wt. % = 48.0 g 75.5 wt. % = 108.3 g                       ______________________________________                                    

The aqueous phase was cooled to T=26° C. The salt crystallisateprecipitated out, consisting of NH₄ HSO₄ +(NH₄)₂ SO₄, was filtered off.The composition obtained is given in Table 4.

                  TABLE 4                                                         ______________________________________                                                Salt crystallisate                                                                         Aqueous filtrate                                           (39.8 g) (103.6 g)                                                          ______________________________________                                        Total MHA 0.1 wt. % (calc.)                                                                            5.0 wt. %                                               = 0.42 g = 5.18 g                                                             = 0.2% of theor. = 2.4% of theor.                                            H.sub.2 O 7.0 wt. % = 2.8 g 25.8 wt. % = 26.7 g                               Salt (calc.) 92.9 wt. % = 37.0 g 69.2 wt. % = 71.72 g                       ______________________________________                                    

The salt crystallisate was washed on the filter with 10 g of MTBE andthe organic filtrate obtained (6.0 g) was analysed (4.8 g loss onevaporation of MTBE):

Total MHA=7.4 wt. %=0.44 g=0.2% of theoretical MHA loss via saltcrystallisate: <0.2% of theoretical without salt wash and 0% oftheoretical with salt wash

The washed salt crystallisate was dried (35.8 g) and analysed:

SO₄ ²⁻ 80.5 wt. %

NH₄ ⁺ 18.5 wt. %

Salt 22.3% of theoretical

The organic phase (292.4 g), the aqueous filtrate (103.6 g) and theorganic filtrate (6.0 g) were mixed with 232 g of MTBE and stirredintensively at room temperature for a brief period. After stirring hadbeen concluded, the two liquid phases formed were separated from oneanother. The phases separated from one another had the compositionsshown in Table 5.

                  TABLE 5                                                         ______________________________________                                                  Organic                                                               extract solution Raffinate                                                    (475 g) (159 g)                                                             ______________________________________                                        Total MHA   46.0 wt. %    1.8 wt. %                                              = 218.5 g (calc.) = 2.86 g                                                    = 98.4% of theor. = 1.3% of theor.                                           H.sub.2 O 3.5 wt. % 25.3 wt. %                                                 = 16.6 g = 40.2 g                                                            NH.sub.4.sup.+ 0.034 wt. % not observed                                        = 0.2 g                                                                      SO.sub.4.sup.2- 0.55 wt. % not observed                                        =2.6 g                                                                       Salt (calc.)  72.9 wt. %                                                        = 115.9 g                                                                 ______________________________________                                    

The residual content of total MHA from the raffinate of the one-stepextraction described here can be decreased to <0.1% of theoretical bysubsequent extraction one or more times using fresh solvent, or bycontinuous extraction in a system comprising several theoretical plates.

The salt crystallisate (35.8 g) was dissolved in the raffinate (159 g)at 61° C. to form a clear solution. The salt solution thus obtained hadthe following composition:

Total MHA 1.4 wt. %=1.3% of theoretical

H₂ O 20.7 wt. %

Salt 77.9 wt. %

The solution thus obtained can with particular advantage be passed to aplant for the recovery of sulphuric acid, as its salt content isdefinitely more than 60 wt. %.

Experiment 6 MHA Separation by Liquid/Liquid Phase Separation andLiquid/Liquid Extraction Using MIBK

505 g of MHA hydrolysate containing 43.7 wt. % (220.7 g) of total MHA(prepared as described in Example 1) was evaporated at a pressure of 50mbar to a content of 49.9 wt. % of total MHA. The concentrate (440 g)consisted of two liquid phases, which were separated from one another atT=60° C. The composition of the two phases is given in Table 6.

                  TABLE 6                                                         ______________________________________                                                   Organic phase                                                                             Aqueous phase                                            (299 g) (141 g)                                                             ______________________________________                                        Total MHA    71.8 wt. %    4.0 wt. %                                             = 214.7 g = 5.6 g                                                             = 97.3% of theor. = 2.5% of theor.                                         ______________________________________                                    

The aqueous phase was cooled to T=20° C. The salt crystallisateprecipitated out, consisting of NH₄ HSO₄ +(NH₄)₂ SO₄, was filtered off.The composition obtained may be seen in Table 7.

                  TABLE 7                                                         ______________________________________                                                   Salt crystallisate                                                                        Aqueous phase                                            (56 g) (83 g)                                                               ______________________________________                                        Total MHA    2.5 wt. % (calc.)                                                                           5.0 wt. %                                             = 1.4 g = 4.15 g                                                              = 0.6% of theor. = 1.9% of theor.                                          ______________________________________                                    

The salt crystallisate was washed on the filter with 14 g of MIBK andthe organic filtrate obtained (13.8 g) was analysed.

Total MHA=9 wt. % 1.2 g=0.56% of theoretical MHA loss via saltcrystallisate: <0.1% of theoretical

The washed salt crystallisate was dried (40 g) and analysed:

SO₄ ²⁻ 80.7 wt. %

NH₄ ⁺ 18.8 wt. %

Salt 24.7% of theoretical

The organic phase (299 g), the aqueous filtrate (83 g) and the organicfiltrate (13.8 g) were mixed with 250 g of MIBK and stirred intensivelyat room temperature for a brief period. After stirring had beenconcluded, the two liquid phases formed were separated from one another.These had the compositions shown in Table 8.

                  TABLE 8                                                         ______________________________________                                                   Organic                                                              extract solution Raffinate                                                    (484 g) (144 g)                                                             ______________________________________                                        Total MHA                  2.05 wt. %                                            217.1 g (calc.) = 2.95 g                                                      = 98.4% of theor. = 1.3% of theor.                                         ______________________________________                                    

The residual content of total MHA from the raffinate of the one-stepextraction described here can be decreased to <0.1% of theoretical bysubsequent extraction one or more times using fresh solvent, or bycontinuous extraction in a system comprising several theoretical plates.

EXAMPLE 5 Recovery of Salt from MHA Hydrolysate Prior to the Extractionby Liquid/Solid Phase Extraction (cf. FIG. 2) Experiment 7 MHASeparation Without Liquid/Liquid Phase Separation

505 g of MHA hydrolysate containing 43.7 wt. % (220.7 g) of total MHA(prepared as described in Example 1) was evaporated at a pressure of 50mbar to a content of 49.9 wt. % of total MHA. The concentrate (440 g)was cooled to room temperature, with a suspension of salt crystallisateand a homogeneous liquid phase being obtained, which was separated byfiltration. The composition found is shown in Table 9.

                  TABLE 9                                                         ______________________________________                                                   Filtrate    Salt crystallisate                                       (342.4 g) (95.0 g)                                                          ______________________________________                                        Total MHA    56.6 wt. %    34.4 wt. % (calc.)                                    = 193.8 g = 32.7 g                                                            = 87.8% of theor. = 14.8% of theor.                                        ______________________________________                                    

The salt crystallisate was washed on the filter with 20 g of MIBK andthe organic filtrate (41.6 g) was analysed: Total MHA: 52.5 wt. %=21.8g=9.9% of theoretical.

The washed salt crystallisate was dried (52.8 g) and analysed:

Total MHA 6.1 wt. %=1.5% of theoretical

SO₄ ²⁻ 75.0 wt. %

NH₄ ⁺ 18.4 wt. %

Salt 93.9 wt. %=30.7% of theoretical

The loss of MHA via the washed salt crystallisate was 1.5% oftheoretical.

The filtrate (342.4 g) was taken up at room temperature in 244 g of MIBKand the organic filtrate (41.6 g) was added thereto, with an aqueousliquid phase separating. The two liquid phases were separated from oneanother and the composition found was that shown in Table 10.

                  TABLE 10                                                        ______________________________________                                                   Organic                                                              extract solution Raffinate                                                    (480 g) (140 g)                                                             ______________________________________                                        Total MHA                  2.25 wt. %                                            212.5 g (calc.) = 3.15 g                                                      = 96.3% of theor. = 1.4% of theor.                                         ______________________________________                                    

The residual content of total MHA from the raffinate of the one-stepextraction described here can be decreased to <0.1% of theoretical bysubsequent extraction one or more times using fresh solvent, or bycontinuous extraction in a system comprising several theoretical plates.

The residual content of total MHA in the salt crystallisate can befurther decreased by additional rewashing with solvent or water. Arewashing with water is preferably carried out using an aqueous solutionof NH₄ HSO₄ and/or (NH₄)₂ SO₄, which is again preferably used severaltimes and, at the latest on becoming completely concentrated, isreturned to the extraction system for reextraction of the total MHAdissolved therein.

The MHA-containing organic filtrate can be returned to the solventextraction system in order to isolate MHA from the organic filtrateand/or from the aqueous filtrate. At the same time a loss of MHA ofabout 0.5 to 12.5% of theoretical is advantageously avoided.

The salt crystallisates from Examples 4 and 5 are suitable forprocessing into marketable (NH₄)₂ SO₄ by the addition of appropriateproportions of NH₃ and subsequent crystallisation. They may also bepassed directly or, preferably, after being dissolved in water or in asuitable solution containing NH₄ HSO₄ or (NH₄)₂ SO₄ or both salts, to aplant for the recovery of H₂ SO₄.

EXAMPLE 6 Liquid/Liquid Extraction of MHA Hydrolysate (cf. FIG. 3)Experiment 8 Extraction Using MTBE

100 g of MHA hydrolysate containing 43.7 wt. % (43.7 g) of total MHA(prepared as described in Example 1) was mixed with 60 g of MTBE at roomtemperature and stirred intensively at room temperature for a briefperiod. After stirring had been concluded, the two liquid phases formedwere separated from one another. The results are compiled in Table 11.

                  TABLE 11                                                        ______________________________________                                                   Organic                                                              extract solution Raffinate                                                    (107 g) (52 g)                                                              ______________________________________                                        Total MHA    39.9 wt. %    2.0 wt. %                                             (43.6 g) (1.0 g)                                                              = 97.7% of theor. = 2.3% of theor.                                           H.sub.2 O 4.16 wt. % 38.7 wt. %                                               NH.sub.4.sup.+ 0.02 wt. % 9.6 wt. %                                           SO.sub.4.sup.2- 0.21 wt. % 50.3 wt. %                                       ______________________________________                                    

Experiment 9 Extraction Using MIBK

100 g of MHA hydrolysate containing 43.7 wt. % of total MHA wasextracted using 60 g of MIBK in a manner similar to Experiment 8 (Table12):

                  TABLE 12                                                        ______________________________________                                                   Organic                                                              extract solution Raffinate                                                    (107.5 g) (51.5 g)                                                          ______________________________________                                        Total MHA    39.0 wt. %    2.4 wt. %                                             (41.9 g) (1.2 g)                                                              = 95.9% of theor. = 2.8% of theor.                                           H.sub.2 O 4.65 wt. % 38.2 wt. %                                               NH.sub.4.sup.+ 0.05 wt. % 9.7 wt. %                                           SO.sub.4.sup.2- 0.38 wt. % 50.4 wt. %                                       ______________________________________                                    

A comparison of the two single-step solvent extractions shows that inthe case of MTBE (Experiment 8) the organic extract solution takes uponly half as much unwanted inorganic ammonium salt as in the case ofMIBK (Experiment 9). Moreover even less MHA is lost via the raffinatephase when MTBE is used.

EXAMPLE 7 MHA Isolation After Increasing the Salt Content (cf. FIG. 4)Experiment 10 MHA Separation by Liquid/Liquid Phase Separation andLiquid/Liquid Extraction

23 g of (NH₄)₂ SO₄ in 598 g of MHA hydrolysate was dissolved in 43.7 wt.% (261.3 g) of total MHA (prepared as described in Example 1) at T=60°C. The solution (621 g) consisted of two liquid phases, which wereseparated at T=60° C. The following composition (Table 13) was found:

                  TABLE 13                                                        ______________________________________                                                   Organic phase                                                                             Aqueous phase                                            (416 g) (205 g)                                                             ______________________________________                                        Total MHA    58.7 wt. %    7.8 wt. %                                             = 244.2 g = 16 g                                                              = 93.4% of theor. = 6.1% of theor.                                           H.sub.2 O 20.4 wt. % 32.7 wt. %                                                = 84.9 g = 67.0 g                                                            Salt (calc.) 20.9 wt. % 59.5 wt. %                                             = 86.9 g = 122.0 g                                                         ______________________________________                                    

The aqueous phase having a salt content of about 60 wt. % can be passeddirectly to a plant for the recovery of sulphuric acid.

The organic filtrate (416 g) was mixed with 250 g of MTBE and stirredintensively at room temperature for a brief period. After stirring hadbeen concluded, the two liquid phases formed were separated from oneanother. Their composition is shown in Table 14.

                  TABLE 14                                                        ______________________________________                                                   Organic                                                              extract solution Raffinate                                                    (522 g) (144 g)                                                             ______________________________________                                        Total MHA    47.5 wt. %    2.4 wt. %                                             = 248.0 g (calc.) = 3.4 g                                                     = 94.9% of theor. = 1.3 of theor.                                            H.sub.2 O 5.5 wt. % 39.0 wt. %                                                 = 28.8 g = 56.2 g                                                            NH.sub.4.sup.+ 0.032 wt. % not observed                                        = 0.2 g                                                                      SO.sub.4.sup.2- 0.28 wt. % not observed                                        = 1.5 g                                                                      Salt -- 58.6 wt. %                                                            (calc.)  = 84.4 g                                                           ______________________________________                                    

The residual content of total MHA from the raffinate of the single-stepextraction described here can be decreased to <0.1% of theoretical bysubsequent extraction one or more times using fresh solvent, or bycontinuous extraction in a system comprising several theoretical plates.

The aqueous phase (205 g) and the raffinate (144 g) were combined. Thesalt solution thus produced (349 g) had the following composition:

Total MHA 5.6 wt. %

H₂ O 35.3 wt. %

Salt 59.1 wt. %

The approx. 60% salt solution can be passed directly to a plant for therecovery of sulphuric acid. A further concentration can be achieved byincreased input of salt in the hydrolysate and by continuous extractionand complete separation of MHA from the raffinate.

Experiment 11 Separation of MHA by Liquid/Liquid Extraction

In a manner similar to Experiment 10, 23 g (NH₄)₂ SO₄ was dissolved in598 g of MHA hydrolysate. 372 g of MTBE was added to the resultingmixture comprising two liquid phases (620 g) and the whole wasintensively stirred at T=40° C. After stirring had been concluded, thetwo liquid phases formed were separated from one another. The resultsare summarised in Table 15.

                  TABLE 15                                                        ______________________________________                                                   Organic                                                              extract solution Raffinate                                                    (658 g) (332 g)                                                             ______________________________________                                        Total MHA    38.8 wt. %    1.9 wt. %                                             = 255.0 g (calc.) = 6.3 g                                                     = 97.6% of theor. = 2.4 of theor.                                            H.sub.2 O 4.4 wt. % 37.0 wt.%                                                  = 30.0 g = 122.8 g                                                           NH.sub.4.sup.+ 0.012 wt.% not observed                                         = 0.08 g                                                                     SO.sub.4.sup.2- 0.15 wt. % not observed                                        0.99 g                                                                       Salt  61.1 wt. %                                                              (calc.)  = 202.9 g                                                          ______________________________________                                    

The residual content of total MHA from the raffinate of the single-stepextraction described here can be decreased to <0.1% of theoretical bysubsequent extraction one or more times using fresh solvent, or bycontinuous extraction in a system comprising several theoretical plates.

The raffinate, having a salt content of >60 wt. %, can be passeddirectly to a plant for the recovery of sulphuric acid.

Washing of the organic extract solution with water can be dispensed withboth in Experiment 10 and in Experiment 11, as the residual sulphatesalt content is already extremely low. This is a great advantage, as inthis way it is possible to avoid both additional operational expense andthe undesirable dilution of the raffinate.

The organic extract solutions produced in Examples 4, 5, 6 and 7,similarly to those in Examples 2 and 8, can be continuously evaporatedto the point of virtually complete removal of the solvent and to a watercontent of <5 wt. %. The MHA high concentrate thus produced can byappropriate conditioning be converted into various MHA product mixtures.

EXAMPLE 8 Description of the Method with Reference to FIG. 6

A diagrammatic arrangement of the apparatus used for Example 8 is shownin FIG. 6. The reference numbers used indicate the following equipment,which substantially constitute the apparatus employed:

(001) extraction column, for example, a pulsed perforated-plate columnof 3 m in length, 2.1 cm internal diameter, having 60 perforated platesand a heated double jacket;

(002) film evaporator, for example, a Sambay evaporator having 0.08 m²exchange surface and a heated double jacket;

(003) condensation system, for example, a water-cooled glass condenser;

(004/005) receiver for returned water or returned solvent

(006) phase separator for organic and aqueous raffinate

(007) washing zone for overflowing extract solution

The MHA hydrolysate resulting from the MHA hydrolysis step, whichconsists substantially of MHA (monomer+dimers+oligomers+optionallyamide), (NH₄)₂ SO₄ and/or NH₄ HSO₄ as well as water, after beingpreheated to the extraction temperature, is introduced into theextraction column 001 above the fortieth plate. The solvent (here methylisobutyl ketone=MIBK) is likewise preheated and pumped into the bottomof the column (countercurrent principle). In addition the overflow fromthe column is subjected to washing water in a washing zone and thewashing phase is returned to the inlet flow of hydrolysate. The aqueousraffinate containing substantially (NH₄)₂ SO₄ and/or NH₄ HSO₄ and waterand the organic raffinate consisting mainly of solvent are withdrawntogether at the bottom of the column, with cooling. The two phases areseparated in a phase separator 006, the organic raffinate is returned tothe extraction system and the aqueous raffinate is transferred out. Theextraction solution containing substantially MHA, solvent and water iswithdrawn at the head of the column and, after being passed through thewashing zone 007, is then fed into the Sambay evaporator 002. There,under a vacuum and additional blowing in of H₂ O vapour as well as of astream of N₂ shortly before the discharge from the evaporator, MIBK andH₂ O are together removed from the extraction solution. The evaporationwas carried out in such a way that <2 wt. % of H₂ O was detectable inthe discharge from the Sambay evaporator and the MHA high concentrateflowing out was virtually free from solvent.

The solvent/water mixture issuing from the evaporator 002 was first ofall condensed in 003 and passed into a separating vessel in order to beseparated. Water and solvent were each collected in a receiver 004 and005 respectively and from there were returned to the extraction system.The discharge from the Sambay evaporator was cooled to room temperatureand passed to a receiver intended for the product.

The composition of the extraction solution was analysed immediatelyafter its leaving the washing zone 007 and the composition of theaqueous and of the organic raffinate solution respectively was analysedin each case immediately after their leaving the phase separator 006.

The composition of the MHA high concentrate was determined in thedischarge from the bottom of the Sambay evaporator immediately after theoutlet point.

The MHA hydrolysate solution used for the extraction was prepared from114.7 kg (874 mol) of MMP cyanohydrin and 131.9 kg (874 mol, 1.00 molequiv.) of 65% H₂ SO₄ in a pressure-resistant 400 l agitated tank at atemperature of 50° C. and with a residence time of 60 min, followed bydilution with 96.7 kg of H₂ O and further reaction at a temperature of90° C. and with a residence time of 120 min. The crude hydrolysatesolution, after conclusion of the reaction, was freed from volatileby-products present by the application of a vacuum and subsequentlyanalysed. The composition thus obtained of the MHA hydrolysate used forthe extraction is given in Example 8, Experiment 12.

The conditions and results of Experiment 12 are summarised in tabularform below.

Experiment 12 Use of MHA Hydrolysate Obtained from MMP Cyanohydrin and1.0 Mol Equiv. of H₂ SO₄

Use in extraction:

Flow rates:

MIBK 6.7 kg/h

MHA hydrolysate 12.3 kg/h

total MHA 4.9 kg/h

washing H₂ O 1.3 kg/h

MIBK hydrolysate 0.55 [-]

Composition of MHA hydrolysate:

total MHA 39.5 wt. %

MHA 94.7 mol %

DIM+OLI: 5.3 mol %

H₂ O: 28.7 wt. %

SO₄ ²⁻ : 27.5 wt. %

(NH₄ HSO₄): 33.2 wt. %

Extraction (001):

Temperature: 60° C. (average)

Compositions

of the extraction solution:

MIBK 44.7 wt. % (calc.)

total MHA 41.8 wt. %

H₂ O 13.5 wt. %

of the aqueous raffinate:

MIBK 77.5 ppm

total MHA 0.1 wt. %

of the organic raffinate:

MIBK 97.5 wt. % (flow rate 0.017 kg/h)

Evaporation 002

Pressure 600 mbar

Sambay

temperature

in the heating jacket 180° C.

at the top 85° C.

at the bottom not observed

stripping steam 0.5 kg/h

stripping gas N₂ 100 l/h

Composition of the MHA high concentrate in the discharge from the bottomof the column:

total MHA 98 wt. %

MHA 86 mol %

DIM+OLI: 14 mol %

H₂ O: 2 wt. %

MIBK 40 ppm

Approx. 4.9 kg/h of MHA high concentrate having the above compositionwas obtained from the discharge from the bottom of the Sambayevaporator. The organic raffinate was returned to the extraction column.The aqueous raffinate was transferred out for disposal directly andwithout further aftertreatment.

It was thus possible to avoid an additional distillation or strippingstep for the removal of residual solvent from the discharge from thebottom of the column. The organic raffinate, which was withdrawn fromthe extraction column under mild conditions as the third liquid phase,could moreover be directly returned to the column without furtherpurification.

This application claims priority from German application 195 48 538 6,filed Dec. 23, 1995, and is the National Phase of PCT/EP96/05437, filedDec. 5, 1996, the entire disclosures of which are incorporated herein byreference.

What is claimed is:
 1. A method for the isolation of2-hydroxy-4-methylthiobutyric acid (MHA), comprising:adding hydrogencyanide (HCN) to methylmercaptopropionaldehyde (MMP); and hydrolysingthe thus formed methylmercaptopropionaldehyde cyanohydrin (MMP-CH)by,(a) in a first hydrolysing step, adding sulphuric acid thereto,thereby forming a reaction mixture containing substantially2-hydroxy-4-methylthiobutyroamide (MHA-amide) and; (b) in a secondhydrolysing step, adding water to the MHA-amide thus formed, therebyforming a reaction mixture containing substantially2-hydroxy-4-methylthiobutyric acid (MHA); bringing the MHA containingreaction mixture into contact in a liquid/liquid extraction system withan organic solvent substantially immiscible with water, thereby formingan extraction solution which contains the solvent and the MHAtransferred out of the reaction mixture, and isolating the MHA as theextract from this extraction solution by evaporation, with the provisothat in the first hydrolysing step (a), MMP-CH is hydrolyzed using from60 to 85% sulphuric acid in the molar ratio of MMP-CH to H₂ SO₄ of from1.0:0.5 to 1:1.0 at a temperature of from 30 to 90° C., and in thesecond hydrolysing step (b) the MHA amide is hydrolyzed by the additionof water without further addition of H₂ SO₄, at a temperature of up to140° C., and an initial salt content of the reaction mixture, prior tothe liquid/liquid extraction, is brought to a concentration of about >50wt. % (wt./wt.), with reference to the sum of the inorganic constituentsof the reaction mixture.
 2. The method according to claim 1, wherein theinitial salt content is adjusted by adding ammonium sulphate to thereaction mixture prior to liquid/liquid extraction.
 3. The methodaccording to claim 2, comprising:prior to isolation of the MHA, addingammonium sulphate in a quantity effective for salting out.
 4. The methodaccording to claim 1, further comprising:increasing the saltconcentration by evaporation.
 5. The method according to claim 1,wherein:at least three liquid phases result from the extraction system.6. The method according to claim 5, comprising forming:a homogeneousextract and a raffinate having first and second liquid phases.
 7. Themethod according to claim 6, wherein:the first liquid phase in theraffinate consists substantially of ammonium salt and water and of smallportions of MHA and organic solvent, and the second liquid phaseconsists substantially of organic solvent and of small portions of waterand MHA.
 8. The method according to claim 7, wherein:the first liquidphase contains water in a quantity of from 20 to 50 wt. %, MHA in aquantity of from 0.01 to 0.5 wt. % and salt in a quantity of from 50 to80 wt. %, and the second liquid phase contains MHA in a quantity of from0.01 to 0.5 wt. %, solvent in a quantity of from 90 to 99.9 wt. %, andwater in a quantity of from 0.1 to 10 wt. %, the constituents of eachphase taken separately totalling 100 wt. %.
 9. The method according toclaim 1, wherein:hydrolysis is carried out in the molar ration of MMP-CHto H₂ SO₄ of from 1:0.6 to 1:0.95.
 10. The method according to claim 1,wherein the organic solvent used to form the extraction solventcomprises one or more ethers.
 11. The method according to claim 10,wherein:the solvent comprises an asymmetrical ether.
 12. The methodaccording to claim 11, wherein:the solvent comprises an ether having aboiling point of <60° C.
 13. The method according to claim 10,wherein:the solvent comprises methyl tertiary butyl ether (MTBE). 14.The method according to claim 1, comprising:evaporating the extractionsolution to a water content of <4%, and distilling off residual solventfrom resulting concentrate.
 15. The method according to claim 14,wherein the step of distilling off the residual solvent comprisesstripping using steam.
 16. The method according to claim 14 or claim 15,wherein,after removal of the solvent a product having a water content of<4% is obtained.
 17. The method according to claim 16, wherein,theproduct has a water content of <3%.
 18. The method according to claim 1,wherein the salt content of the reaction mixture is brought to aconcentration of >55 wt. % (wt./wt.).